Process for the production of ethers, typically thf

ABSTRACT

A process is described for the production of ethers, typically terahydrofuran, by reaction of a corresponding organic feed material selected from dicarboxylic acids and/or anhydrides, monoesters of dicarboxylic acids and/or anhydrides, diesters of dicarboxylic acids and/or anhydrides, lactones, and mixtures of two or more thereof in the presence of hydrogen which comprises the steps of: (a) supplying a stream comprising the organic feed material to a first vaporisation zone and contacting said feed with cycle gas comprising hydrogen such that at least a portion of the feed material is vaporised by and into the cycle gas; (b) supplying at least a portion of the cycle gas and the vaporised feed material to a first reaction zone comprising catalyst and operating under reaction conditions to allow hydrogenation and dehydration to occur; (c) recovering from the first reaction zone an intermediate product stream comprising unreacted feed material, cycle gas, desired product(s), and any co-products and by-products; (d) supplying the intermediate product stream to a second vaporisation zone and contacting it with additional feed material such that the said additional feed material is vaporised by and into the intermediate product stream; (e) supplying the product of step (d) to a subsequent reaction zone comprising catalyst and operating under reaction conditions to allow hydrogenation and, if required, dehydration to occur; and (f) recovering from the subsequent reaction zone a product stream comprising the ether.

[0001] The present invention relates to the production of ethers,optionally with the co-production of diols and/or lactones by reactionof an organic feed material in the presence of hydrogen. The reactionwill generally be by hydrogenation and/or dehydration. The organic feedmaterial is selected from dicarboxylic acids and/or anhydrides,monoesters of dicarboxylic acids and/or anhydrides, diesters ofdicarboxylic acids and/or anhydrides, lactones, a mixture thereof or amixture of two or more thereof. In particular it relates to theproduction of C₄ to C₁₂ ethers, optionally with the co-production of thecorresponding diols and/or lactones by the reaction of di-(C₁ toC₄)alkyl esters of C₄ to C₁₂ dicarboxylic acids and/or anhydrides in thepresence of hydrogen. More particularly, it relates to the production ofcyclic ethers.

[0002] More particularly, the present invention relates to a process forthe co-production of C₄ compounds, more specifically tetrahydrofuran,butane-1,4-diol and/or γ-butyrolactone from a hydrocarbon feedstockcomprising a dialkyl maleate by vapour phase reaction in a hydrogen richstream. In a particularly preferred arrangement of the presentinvention, it relates to a process for the production of at least 20% oftetrahydrofuran with co-production of butane-1,4-diol and/orγ-butyrolactone. In the most preferred arrangement it relates to theproduction of tetrahydrofuran with any residual butane-1,4-diol and/orγ-butyrolactone being recycled and converted to further tetrahydrofuran.

[0003] It is known to produce diols by hydrogenation of dialkyl estersof dicarboxylic acids and/or anhydrides, lactones, and mixtures thereofwith a minor amount, typically no more than about 10 wt/wt % andpreferably no more than 1 wt/wt %, of a monoester of the dicarboxylicacid and/or anhydride. Commercial plants have been built which producebutane-1,4-diol as the primary product with small amounts, typically upto about 10 mole %, of tetrahydrofuran and up to about 15 mole % ofγ-butyrolactone by hydrogenation of a dialkyl ester of maleic acidand/or anhydride, such as dimethyl maleate or diethyl maleate, which maycontain minor amounts of dialkyl fumarate and/or dialkyl succinate.Dimethyl succinate or diethyl succinate have also been suggested assuitable starting materials for hydrogenation to producebutane-1,4-diol, tetrahydrofuran and γ-butyrolactone. These succinatesmay be formed by any suitable manner and may be from biotechnologysources.

[0004] For further information regarding the operation of these plantsreference may be made, for example, to U.S. Pat. No. 4,584,419, U.S.Pat. No. 4,751,334, WO-A-86/03189, WO-A-88/00937, U.S. Pat. No.4,767,869, U.S. Pat. No. 4,945,173, U.S. Pat. No. 4,919,765, U.S. Pat.No. 5,254,758, U.S. Pat. No. 5,310,954 and WO-A-91/01960, the disclosureof each of which is herein incorporated by reference.

[0005] Whilst many plant operators aim to maximise the yield ofbutane-1,4-diol and to minimise the yield of the co-products,tetrahydrofuran and γ-butyrolactone, these co-products are themselvesvaluable commodity chemicals. The tetrahydrofuran is normally recoveredas it is an important monomer for making elastomer fibres and is also animportant solvent and therefore is a commercially important chemical.The γ-butyrolactone may be recovered but, as the market for this productis small, it is often recycled to the hydrogenation step for conversionto further butane-1,4-diol and the co-product tetrahydrofuran.

[0006] The dialkyl maleates which are used as feedstock in suchhydrogenation processes may be produced by any suitable means. Thehydrogenation of dialkyl maleates to yield butane-1,4-diol is discussedin detail in U.S. Pat. No. 4,584,419, U.S. Pat. No. 4,751,334 andWO-A-88/00937, which are incorporated herein by reference.

[0007] One conventional process for the production of butane-1,4-dioland co-product tetrahydrofuran with optional production ofγ-butyrolactone is illustrated schematically in FIG. 1. In this process,a dialkyl ester, such as dimethyl maleate together with any residualmethanol from the esterification reactor, is fed via line 1 to avaporiser 2 where it is vaporised into a stream of hot cycle gas whichis usually pre-heated. Cycle gas will normally contain a highconcentration of hydrogen gas but may also include other gases includinghydrocarbons, carbon oxides, methane, nitrogen. Further, where the cyclegas includes recycled gases from downstream, condensables includingproduct ether, methanol, water, co-products and by-products may also bepresent.

[0008] The cycle gas is fed to the vaporiser 2 in line 3. The combinedvaporous stream is then passed in line 4 to the reactor 5 where it isreacted to form butane-1,4-diol, tetrahydrofuran and/or γ-butyrolactone.The product stream 6 is cooled and the reaction products are condensedat 7 and separated from the cycle gas before being passed in line 8 to arefining zone 9. Recovered cycle gas is compressed and recycled in line10. Make-up hydrogen will be added to the recovered cycle gas in line 11with the enriched cycle gas being fed back to vaporiser 2. In therefining zone 9 the various products are separated and thebutane-1,4-diol is removed in line 12 and the tetrahydrofuran in line13. The γ-butyrolactone, together with the intermediate dimethylsuccinate and some butane-1,4-diol may be recycled in lines 14 and 15.In one arrangement the γ-butyrolactone may be partially extracted in anoptional refining zone 16 and removed in line 17. The methanol waterstream separated from the product mix will be recycled upstream via line18.

[0009] A significant portion of the butane-1,4-diol produced by this orother conventional methods is subsequently converted to tetrahydrofuran.This conversion step has substantial cost implications both ininvestment and operation of the plant required for the conversion and asthe importance of tetrahydrofuran increases together with its use inderivative applications, it is desirable to provide a process for theproduction of tetrahydrofuran without the need for this expensivedownstream processing. The downstream processing of conventional methodsincludes recovering the butane-1,4-diol, reacting it to form thetetrahydrofuran and then refining the tetrahydrofuran product.

[0010] In conventional processes, the quantity of cycle gas required tovaporise the feed is determined by a number of parameters including theoperating pressure, the desired reaction temperature, the vaporiser exittemperature and the vapour pressure of the components to be vaporised.

[0011] Whilst it may be desirable to minimise the amount of cycle gasrequired, with prior art systems, this decrease will require the exittemperature from the vaporiser to be maintained high. However,maintaining a high vaporisation exit temperature would mean that thereaction temperature would be higher than desired. It is desirable tomaintain the operating temperature as low as possible for severalreasons including avoidance of hydrogen embrittlement of carbon steelequipment, avoidance of excessive catalyst deactivation and to minimisethe formation of by-products such as butanol.

[0012] It will therefore be understood that the amount of cycle gasrequired for the reaction is determined by the vaporiser exittemperature and is therefore a compromise between the high temperaturenecessary to minimise the cycle gas required to vaporise the feed andthe relatively low temperatures required for the reasons given above.

[0013] In the particular prior art system of the type illustrated inFIG. 1, in which the butane-1,4-diol is the main product, at a reactorinlet temperature of about 165° C. and a pressure of about 63 barapproximately 240 moles of cycle gas are required per mole of dimethylmaleate to be vaporised. Although the temperature will rise across thereactor, the reactor outlet stream will have about the same degree ofsaturation as the inlet stream because the vapour pressure of thebutane-1,4-diol is less than that of the dimethyl maleate in the feed.Since the byproduct γ-butyrolactone and intermediate dimethyl succinate,together with the associated butane-1,4-diol are conventionally recycledto the reaction system, additional cycle gas is required to vaporise therecycle stream(s). This will typically increase the cycle gasrequirements to about 310 mols of cycle gas per mole of the dimethylmaleate vaporised, which it will be understood is a significantincrease.

[0014] Typically a process of the type illustrated in FIG. 1 willproduce up to approximately 10 mole % tetrahydrofuran.

[0015] It is therefore desirable to provide a process for the productionof higher mole % of tetrahydrofuran without the need for expensivedownstream processing. It is further desirable to provide a process inwhich the cycle gas requirements are minimised such that investment andoperating costs are reduced as the selectivity to tetrahydrofuran isincreased. Thus according to the present invention there is provided aprocess for the production of an ether by reaction of a correspondingorganic feed material selected from dicarboxylic acids and/oranhydrides, monoesters of dicarboxylic acids and/or anhydrides, diestersof dicarboxylic acids and/or anhydrides, lactones, and mixtures of twoor more thereof in the presence of hydrogen which comprises the stepsof:

[0016] (a) supplying a stream comprising the organic feed material to afirst vaporisation zone and contacting said feed with cycle gascomprising hydrogen such that at least a portion of the feed material isvaporised by and into the cycle gas;

[0017] (b) supplying at least a portion of the cycle gas and thevaporised feed material to a first reaction zone comprising catalyst andoperating under reaction conditions to allow hydrogenation anddehydration to occur;

[0018] (c) recovering from the first reaction zone an intermediateproduct stream comprising unreacted feed material, cycle gas, desiredproduct(s), and any co-products and by-products;

[0019] (d) supplying the intermediate product stream to a secondvaporisation zone and contacting it with additional feed material suchthat the said additional feed material is vaporised by and into theintermediate product stream;

[0020] (e) supplying the product of step (d) to a subsequent reactionzone comprising catalyst and operating under reaction conditions toallow hydrogenation and, if required, dehydration to occur; and

[0021] (f) recovering from the subsequent reaction zone a product streamcomprising the ether.

[0022] In the ether production reaction of the present invention, theconversion of the acid, anhydride and/or the lactone or ester to formthe diol is an ester hydrogenation or hydrogenolysis and the reaction ofthe diol to the ether, is a dehydration reaction.

[0023] Without wishing to be bound by any theory, it is believed thatthe process of the present invention allows that the amount of productproduced as light boiling (higher vapour pressure) ether rather thandiol is increased, such that the outlet dewpoint of the reactor movesbelow the operating temperature such that further feed material can bevaporised into the stream until the stream approaches saturation. Thisis in marked contrast to conventional processes for the production ofdiols which the inlet and outlet of the reactor are close to the vapourdewpoint. The additional feed material vaporised by the process of thepresent invention may then be collected to product in the secondreaction zone. By this means more feed may be processed to product thanwould have been possible with the conventional process unless the gascirculation rate was increased. In this connection it will be understoodthat a key factor in the cost of conventional processes relates to thehydrogenation loops which are themselves dependent on the amount of gasrequired to vaporise the feed; thus an increase of the gas circulationrate is particularly disadvantageous.

[0024] The cycle gas will normally contain a high concentration ofhydrogen gas but may also include other gases including hydrocarbons,carbon oxides, methane, nitrogen. Further, where the cycle gas includesrecycled gases from downstream, condensables including product ether, C₁to C₄ alkanol, water, co-products and by-products may also be present

[0025] In a particularly preferred embodiment of the present inventionthe ether is a cyclic ether. Most preferably the cyclic ether istetrahydrofuran. In this latter case the organic feed material ispreferably dialkyl maleate. Co-products which may be present to agreater or lesser extent in this embodiment or which may be absentinclude butane-1,4-diol and γ-butyrolactone. This reaction isillustrated in Scheme 1. In this example the alkanol is methanol and theintermediate material is partially hydrogenated dimethyl succinate.

[0026] By-products may include the alkanol used in the esterification ofthe acid or anhydride, for example methanol, undesirable material formedin side reactions, for example butanol, water evolved in the dehydrationof the diol to the ether and intermediate material, for example dimethylsuccinate together with other light or heavy materials formed in theprocess.

[0027] The by-products may be separated from the ether in a refiningzone and may be further purified if required. Similarly, the co-productsmay be separated from the ether in the refining zone and may be furtherpurified if required.

[0028] However, in one arrangement, one or more of the co-productsand/or by-products will be recycled to the first vaporisation zone wherethey will be vaporised. In one alternative arrangement, one or more ofthe co-products and/or by-products will be recycled to the secondvaporisation zone where they will be vaporised into the intermediateproduct stream exiting from the first reaction zone.

[0029] Thus, in the preferred embodiment any dialkyl succinate presentas a by-product may be recycled to one of the vaporisers, preferably thesecond vaporiser and hence to the corresponding reaction zone to improvethe overall selectivity of the reaction to the desired tetrahydrofuranand co-products butane-1,4-diol and/or γ-butyrolactone.

[0030] The vapour pressure of tetrahydrofuran (8284 mmHg at 165° C.) issubstantially higher than that of the butane-1,4-diol (76 mmHg at 165°C.), γ-butyrolactone (252 mmHg at 165° C.) and dimethyl maleate (262mmHg at 165° C.). Thus, in the embodiment where tetrahydrofuran isproduced, optionally with butane-1,4-diol and γ-butyrolactone, fromdimethyl maleate as the conversion of feed dimethyl maleate totetrahydrofuran is increased, the dew point of the exit stream from thefirst reaction zone is reduced. This allows for the additional feedand/or the or each optional recycle stream to be vaporised into theintermediate product stream from the first reaction zone.

[0031] Thus as the amount of ether, for example tetrahydrofuran, presentin the intermediate product stream increases with improved selectivity,the capacity of the intermediate product stream to carry the additionalorganic feed, for example dimethyl maleate, and/or the recycle stream asa vapour is increased.

[0032] In the preferred embodiment of the present invention wheredimethyl maleate is used in the formation of tetrahydrofuran, the cyclegas requirement is about 210 mols per mole of dimethyl maleate feed tothe first reaction zone and additional cycle gas is not required tovaporise the recycle stream. Thus, if, for example, the catalyst in thefirst reaction zone gives approximately 50% selectivity totetrahydrofuran then the total cycle gas required to vaporise both thefeed and the recycle stream is reduced from about 310 moles required inthe prior art process of FIG. 1 to about 210 moles per mole of dimethylmaleate. If in the preferred embodiment of approaching 100% selectivityfor tetrahydrofuran is achieved, then about 100 to 110% more dimethylmaleate may be vaporised per mol of the cycle gas circulated comparedwith that achieved in the prior art process of FIG. 1, bringing thecycle gas requirement down to about 150 moles per mole of dimethylmaleate.

[0033] The capital cost of equipment and the operating cost of thereaction process, particularly energy and other utility requirements islargely determined by the cycle gas flow rate in the system. Forexample, the compressor, heat exchangers and interconnecting pipeworkare largely sized on the cycle gas flow rate and the power ofcompression and heat added to, and removed from, the reaction system arelargely determined by the cycle gas flow. Thus increasing the conversionrate of the ester to ether allows that more moles of feed can bevaporised and hence more product made per mole of gas circulated whichwill have the advantage of substantially decreasing the capital andoperating costs.

[0034] Where all of the co-products, such as butane-1,4-diol andγ-butyrolactone, are recycled to the subsequent vaporisation zone toprovide high, preferably total, conversion to the ether e.g.tetrahydrofuran, it will not be necessary to minimise the lactone, forexample γ-butyrolactone, to diol, for example butane-1,4-diol, ratio inthe reaction by operating at high pressure as is required inconventional processes for co-producing butane-1,4-diol as the mainproduct, with tetrahydrofuran and a minor amount of γ-butyrolactone.Indeed, it may be desirable to operate at a lower reaction pressure andhence, higher γ-butyrolactone to butane-1,4-diol ratio than has beendesirable heretofore. This is in part because the γ-butyrolactone has ahigher vapour pressure than the butane-1,4-diol and therefore requiresless moles of cycle gas for vaporisation, but more significantly theinvestment and operating costs are reduced as the reaction pressure islowered.

[0035] The feed material to the, or each, vaporisation zone may be, ormay include, one or more recycle streams. Fresh organic feed andrefining recycle streams may be vaporised together or may be vaporisedin separate parts of the or each vaporisation zone. This is particularlyadvantageous as it will minimise the risk of transesterification betweenthe ester and the diol. In one arrangement, all of the cycle gas and theorganic feed fed to the first vaporisation zone (step a) is supplied tothe first reaction zone (step b) with the remaining organic feed andrefining recycles being vaporised (step d) into the intermediate productstream recovered from the first reaction zone (step c) to form theintermediate feed stream which is fed to the subsequent reaction zone(step d).

[0036] In a second alternative arrangement, the gaseous stream from thefirst vaporiser (step a) may be divided with a major portion, preferablyfrom about 70% to about 80%, being supplied to the first reaction zone(step b) and a minor portion, preferably from about 20% to about 30%,by-passing the first reaction zone and being fed to the subsequentvaporisation zone, preferably one part of the subsequent vaporisationzone (step d), where it is further heated such that additional organicfeed material can be vaporised into the cycle gas before yielding a hotsecondary feed stream. Where the minor portion is fed to one part of thesubsequent vaporisation zone, the intermediate product stream recoveredfrom the first reaction zone (step c) is fed to a second part of thesubsequent vaporisation zone (step d) into which the refining recyclesare fed. The two streams from the two separate parts of the subseqentvaporisation zone are then mixed to yield the intermediate feed streamwhich is fed to the subsequent reaction zone (step e).

[0037] One advantage of this preferred embodiment is that the liquidadditional organic feed, which may be or include an ester, is separatefrom the liquid refining recycles which contain diols and/or lactones,and is only mixed therewith in the vapour phase. This will minimise thecontact time and hence the potential for transesterification andprogressive chain length growth.

[0038] The feed material fed to the or each vaporisation zone may bewholly, or may include, one or more recycle streams

[0039] Whilst the present invention has been described with particularreference to two reaction zones, in one arrangement of the presentinvention, the process includes more than two reaction zones. Wherethere are more than two reaction zones, corresponding vaporisation zonesmay be located between adjacent reaction zones. Vaporisation in thesesubsequent zones may be made directly into the intermediate productstream from the previous reaction zone or if required a supplementarystream of cycle gas which may comprise one or more of fresh organicfeed, refining recycle material and hydrogen may be included. Theorganic feed recycle material and/or hydrogen if present may be heated.

[0040] The organic feed material is preferably selected from mono C₁ toC₄ alkyl esters of C₄ to C₁₂ dicarboxylic acids and/or anhydrides, di C₁to C₄ alkyl esters of C₄ to C₁₂ dicarboxylic acids and/or anhydrides,lactones of C₄ to C₁₂ hydroxycarboxylic acids, and mixtures of two ormore thereof.

[0041] For example, the organic feed material can be selected from monoC₁ to C₄ alkyl esters of C₄ dicarboxylic acids and/or anhydrides, di C₁to C₄ alkyl esters of C₄ dicarboxylic acids and/or anhydrides,γ-butyrolactone, and mixtures of two or more thereof. A particularlypreferred organic feed material may be selected from monomethyl maleate,monomethyl fumarate, monomethyl succinate, dimethyl maleate, dimethylfumarate, dimethyl succinate, γ-butyrolactone, recycle γ-butyrolactoneand/or butane-1,4-diol and mixtures of two or more thereof.Alternatively the organic feed material can be selected from monoethylmaleate, monoethyl fumarate, monoethyl succinate, diethyl maleate,diethyl fumarate, diethyl succinate, γ-butyrolactone, recycleγ-butyrolactone and/or butane-1,4-diol and mixtures of two or morethereof.

[0042] In one arrangement, the organic feed material fed to one or moreof the vaporisation zones is contained within an organic solvent. Wherethe organic solvent is present, one or more of the vaporisation zones isoperated such that the organic feed material is essentially separatedfrom the organic solvent by cycle gas stripping.

[0043] Suitable organic solvents include: di-(C₁ to C₄ alkyl) esters ofalkyl dicarboxylic acids containing up to 13 carbon atoms; mono- anddi-(C₁₀ to C₁₈ alkyl)esters of maleic acid, fumaric acid, succinic acidand mixtures thereof; (C₁ to C₄ alkyl)esters of napthalenemonocarboxylicacids; tri-(C₁ to C₄ alkyl)esters of aromatic tricarboxylic acids;di-(C₁ to C₄ alkyl)esters of isophthalic acid; alkyl phthalates; anddimethyl sebecate.

[0044] The vaporous feed stream to the first reaction zone preferablyhas a hydrogen-containing cycle gas:condensable material molar ratio inthe range of from about 50:1 to about 1000:1.

[0045] Typically the feed temperature to the first hydrogenation zone isfrom about 100° C. to about 300° C., more preferably from about 150° C.to about 250° C., while the feed pressure to the first reaction zone istypically from about 50 psia (about 346 kPa) to about 2000 psia (about13790 kPa), for example, more preferably from about 450 psia (about 3103kPa) to about 1000 psia (about 6895 kPa).

[0046] The hydrogenatable material is preferably supplied to the firstreaction zone at a rate corresponding to a liquid hourly space velocityof from about 0.05 to about 5.0 h⁻¹.

[0047] If desired, the pressure and/or the temperature can be adjustedin any convenient manner between the first and subsequent reaction zonesand/or between adjacent reaction zones where more than two reactionzones are present. The temperature may be adjusted by any suitable meansincluding the use of a heat exchanger or exchangers.

[0048] The hydrogen make up gas used in the process of the presentinvention can be obtained by any conventional manner. Preferably itcontains at least about 50 volume % up to about 99.99 volume % or more,e.g. from about 80 to about 99.9 volume %, of hydrogen. It may furthercontain one or more inert gases, such as nitrogen or methane.Conveniently the hydrogen make up gas is produced by pressure swingabsorption so that the cycle gas molecular weight is minimised therebyreducing the power required for compression and circulation of the cyclegas.

[0049] Any suitable catalyst for the reaction may be selected. Whilst amixture of catalysts may be used, for ease of reference the term“catalyst” will be used herein and will be understood to mean either asingle catalyst or a mixture of two or more different catalysts. Thecatalyst used in the subsequent reaction zone may be different from thatused in the first reaction zone. Where there are more than two reactionzones present, the catalyst used in the or each zone may be the same asor different from that used in the first and/or subsequent reactionzone.

[0050] In one arrangement, a bed comprising a variety of catalysts maybe used. In one example, the bed may include a catalyst that is tolerantof residual feed acid content, one which is suitable to promotehydrogenation of the ester and another which promotes selectivity to thedesired ether. Catalyst beds comprising more than one type of catalystmay comprise discrete layers of catalyst within the bed such thatdifferent types are separated or the different catalyst types may beadmixed.

[0051] In a particularly preferred process the catalyst of the firstreaction zone is selected from noble metal and/or copper-containingcatalysts. Hence the catalyst of the first hydrogenation zone can be orinclude one or more of a palladium catalyst, a reduced copper chromitecatalyst or a reduced copper containing catalyst. The same or adifferent catalyst may also be used in the subsequent and any additionalreaction zones. In one arrangement, the catalyst in at least thesubsequent reaction zone is, or includes, a copper-containing catalyst.

[0052] Examples of copper-containing catalysts include reduced copperoxide/zinc oxide catalysts, reduced manganese promoted copper catalysts,reduced copper chromite catalysts, and reduced promoted copper chromitecatalysts.

[0053] One alternative catalyst for use in at least one of the reactionzones is a reduced manganese promoted copper catalyst.

[0054] When the or each catalyst is a copper-containing catalyst, theactive catalytic species may be at least partially supported on asupporting material selected from chromia, zinc oxide, alumina, silica,silica-alumina, silicon carbide, zirconia, titania, carbon, or a mixtureof two or more thereof, for example, a mixture of chromia and carbon.

[0055] In one preferred process of the present invention an acidtolerant catalyst such as a promoted copper chromite catalyst may beused in at least one of the reaction zones. A suitable promoted copperchromite catalyst is, for example, the catalyst sold as PG85/1 by DavyProcess Technology Limited of The Technology Centre, Princeton Drive,Thornaby, Stockton-on-Tees, TS17 6PY, England.

[0056] A catalyst which is effective to hydrogenate the ester to diolsand lactones such as a manganese promoted copper catalyst may also beused in at least one of the reaction zones. A suitable manganesepromoted copper catalyst which exhibits superior conversion of a dialkylester under typical operating conditions used for catalyst PG85/1 issold as DRD92/89A by Davy Process Technology Limited. A catalyst with ahigh selectivity to the desired ether under typical operating conditionsis DRD92/89B which is also available from Davy Process TechnologyLimited.

[0057] Further details of suitable catalysts can be found inInternational Patent Application No. PCT/GB00/04758 which isincorporated herein by reference.

[0058] Typically the hydrogenatable material will contain from about0.01 to about 1.0 wt/wt % or more, e.g. up to about 10 wt/wt %, butnormally no more than about 2.0 wt/wt/o, of acidic material.

[0059] The charge of catalyst in the first reaction zone is preferablysufficiently large to reduce the content of acidic material to less thanabout 0.005 wt/wt % in passage of the vaporous mixture therethrough.

[0060] The amount of catalyst used in each reaction zone may be the sameor different. The catalyst charge in the first reaction zone mayconstitute from about 10% to about 70%, more usually about 20% to about50%, of the total catalyst volume in the reaction zones. Similarly thecatalyst of the subsequent reaction zone is typically in the range offrom about 70% to about 10%, more usually about 20% to about 50%, of thetotal catalyst volume of the reaction zones.

[0061] The selected catalyst preferably converts the ester, preferablythe dialkyl maleate, to the desired ether, preferably a cyclic ethermost preferably tetrahydrofuran, at a selectivity of from about 20% toabout 90% or more, most preferably, about 70% or more.

[0062] The product stream from the final reaction zone is preferablyfed, preferably having been condensed, to a refining zone where thedesired ether, preferably tetrahydrofuran, is separated as product. Anyco-products, such as butane-1,4-diol and/or γ-butyrolactone, which maybe present may be separated or may be recycled to the reaction system.Where there is more than one co-product, one or more may be separatedand recovered and the remainder recycled.

[0063] In one arrangement where 100% conversion to ether, for exampletetrahydrofuran, is desired all of the co-products, for examplebutane-1,4-diol and/or γ-butyrolactone, are recycled.

[0064] The ability to select suitable catalysts and adjust the recyclingof co-products to the or each vaporisation zone allows the plantoperator flexibility to select the amount of ether produced relative tothe formation of the or each co-product.

[0065] Any alkanol derived from the organic feed, which will typicallybe a C₁ to C₄ alkanol and water in the crude product stream willpreferably be condensed and separated in refining. The alkanol willconventionally be recycled to the esterification reactor in which theorganic feed material is formed, if present. The refining system mayinclude means, if required to separate the water from the alkanol. Therefining system will usually include means to separate other by-productswhich may be recycled. An example of a by-product which may be recycledis an for example any intermediate material. Alternatively some or allof any by-products produced may be rejected as effluent. An example of aby-product which may be rejected is any mono-ol produced.

[0066] The present invention will now be described, by way of example,with reference to the accompanying drawings in which:

[0067]FIG. 1 is a schematic diagram of a prior art arrangement; and

[0068]FIG. 2 is a schematic diagram of a process in accordance with thepresent invention.

[0069] It will be understood by those skilled in the art that thedrawings are diagrammatic and that further items of equipment such asreflux drums, pumps, vacuum pumps, compressors, gas recycle compressors,temperature sensors, pressure sensors, pressure relief valves, controlvalves, flow controllers, level controllers, holding tanks, storagetanks, and the like may be required in a commercial plant. The provisionof such ancillary items of equipment forms no part of the presentinvention and is in accordance with conventional chemical engineeringpractice.

[0070] Whilst for convenience, the description and drawing impliesseparate heat exchange, vaporisation and reaction equipment, it will beunderstood that some or all of these may be included into a singlevessel or each associated vaporisation zone and reaction zone may becontained within a single vessel.

[0071] The present invention will now be described with particularreference to the production of tetrahydrofuran by reaction of a feed ofdimethyl maleate with hydrogen.

[0072]FIG. 2 illustrates a plant for the production of tetrahydrofuranby reaction of dimethyl maleate with hydrogen in the vapour phase. Thedimethyl maleate may be produced by any suitable means and may besupplied from an esterification plant (not shown) of the type describedin WO-A-90/08127 which is incorporated herein by reference.

[0073] The resulting dimethyl maleate typically contains no more thanabout 10.0 wt/wt % of acidic organic materials, such as monomethylmaleate, and preferably less than about 2.0 wt/wt %, e.g. about 0.1 toabout 1.0 wt/wt %, of acidic materials. The dimethyl maleate is fed inline 19 with a portion going to a first vaporisation zone 20 which maycontain packing. The feed may be pumped to near the top of thevaporisation zone. The vaporisation zone is operated at a temperature ofabout 167° C. and a pressure of 900 psia (6205 kPa).

[0074] The feed flows down the vaporisation zone against an upflowingstream of cycle gas from line 21 which may include fresh make uphydrogen fed from line 22 that has been added to recovered cycle gas(line 23) from downstream. Alternatively, it may simply be the recoveredcycle gas from line 23 with the makeup hydrogen may be added elsewherein the system as convenient.

[0075] Where wetting of the catalyst may cause the catalyst todeteriorate it may be desirable to feed the reaction mixture to thereactor above the dew point. This can be achieved by either passing asuitable excess cycle gas flow through the vaporiser or adding extracycle gas flow after the vaporiser, or adding extra heat to the reactionmixture before feeding to the reaction zone. However, if wetting of thecatalyst is not deleterious to the operation of the catalyst, entrainedliquid may be present. The reaction will, however, still be essentiallya vapour phase reaction.

[0076] A near saturated vapour mixture stream comprising dimethylmaleate in cycle gas, with a cycle gas:dimethyl maleate molar ratio ofabout 150:1 is recovered from the top of the vaporisation zone.

[0077] The mixture of gases is then fed in line 24 to the first reactionzone 25 which contains a fixed bed catalyst charge.

[0078] The catalyst charge preferably contains acid tolerant catalystsuch as PG85/1 and DRD92/89A which promote ester hydrogenation and DRD92/89B which promotes diol dehydration. The reaction zone is generallyoperated at an inlet temperature of about 167° C. to about 175° C., aninlet pressure of about 900 psia (6205 kPa), and an exit temperature ofabout 195° C. The dimethyl maleate feed rate corresponds to a liquidhourly space velocity of 0.5 h⁻¹. Partial conversion of dimethyl maleateto butane-1,4-diol, tetrahydrofuran and γ-butyrolactone, as well assmall quantities of undesirable by-products, such as butanol and/oracetal 2-(4′-hydroxybutoxy)-tetrahydrofuran, occurs in passage throughreactor 25. In addition, partial hydrogenation of feed dimethyl maleateto dimethyl succinate occurs. The resulting first intermediate reactionmixture, passes through line 26 into the second vaporisation zone 27.

[0079] Fresh feed is added via line 28 and is mixed with theintermediate reaction mixture into which the fresh feed is vaporised. Itmay also be mixed with one or more recycled refining streams fromdownstream which are added in line 29. The hot intermediate reactionmixture will also vaporise the majority of the recycled material.

[0080] The mixture from vaporisation zone 27 is passed in line 30 to thesecond reaction zone 31, which contains a further charge of catalyst.

[0081] Here the further reaction is carried out and the amount oftetrahydrofuran in the product stream is increased. The product stream32 is passed to a cooler and condenser 33 where the crude product isseparated from the cycle gas which is recycled via a line 34 to acompressor 35 and lines 23 and 21 to the first vaporiser 20.

[0082] Crude product is passed in line 36 to a refining system 37. Herethe crude product stream is separated, preferably by distillation inseveral stages, to yield pure tetrahydrofuran which is recovered in line38. Lines 39 and 40 for the separate recovery of the butane-1,4-diol andthe γ-butyrolactone may be provided or in a preferred arrangement, oneor both of these, optionally together with partially hydrogenated feedmaterial may be recycled in line 29 to the second vaporisation zone forfurther reaction to yield tetrahydrofuran.

[0083] Methanol and water may be recycled to upstream reactors in line41 or may be separated and the methanol recycled in line 42 and thewater extracted as effluent in line 43.

[0084] The invention will now be further described with reference to theaccompanying examples.

COMPARATIVE EXAMPLE 1

[0085] In a prior art process as illustrated in FIG. 1, in order tovaporise 1 kmol/h of which is fed to the vaporiser, 0.4 kmol/h ofrefining recycle, 311 kmol/h of hydrogen cycle gas and 4.9 kmol/h ofmake up hydrogen are also fed to the vaporiser. The vaporised stream isthen fed via line 4 to the reactor where the dimethyl maleate andrefining recycles are converted to crude reaction products. These arecooled and separated and the crude product is fed to a refining zonewhere the products are refined and the refining recycles are recycled tothe vaporiser. The selectivity to tetrahydrofuran is measured. Thereaction details and results are set out in Table 1.

EXAMPLE 1

[0086] In a process scheme in accordance with the present invention andas illustrated in FIG. 2, the compressor cycle gas stream is maintainedat the same absolute rate as that for Comparative Example 1. In thisarrangement, 1.5 kmol/h of dimethyl maleate is fed to the firstvaporiser. No feed dimethyl maleate is fed to the second vaporiser.However, 0.3 kmol/h of refining recycle is fed to the second vaporiser.311 kmol/h of hydrogen cycle gas and 7.6 kmol/h of make up hydrogen arefed to the first vaporiser to vaporise the dimethyl maleate feed and thevaporised stream passes to the first reactor where conversion to crudeproduct occurs. The reactor contains a sufficient quantity of suitablecatalyst to convert approximately 50% of the dimethyl maleate totetrahydrofuran. The stream from this reactor passes to the secondvaporiser where it is used to vaporise refining recycle. The stream fromthe second vaporiser passes to the second reactor where conversion tocrude product occurs. The products from the second reactor are refinedand the refining recycles separated and recycled to the secondvaporiser. The selectivity to tetrahydrofuran is measured. The reactiondetails and results are set out in Table 1. It can be seen thatapproximately 50% more dimethyl maleate is reacted than is possible withthe procedure of the prior art.

EXAMPLE 2

[0087] In a process scheme in accordance with the present invention andas illustrated in FIG. 2, the compressor cycle gas stream is maintainedat the same absolute rate as that for Comparative Example 1. In thisarrangement, a total dimethyl maleate feed of 1.9 kmol/h is fed to thesystem with 1.5 kmol/h being provided to the first vaporiser and 0.4kmol/h to the second vaporiser. 311 kmol/h of hydrogen cycle gas and 9.4kmol/h of make up hydrogen are fed to the first vaporiser to vaporisethe dimethyl maleate feed before the vaporised stream is passed to thefirst reactor where conversion to crude product occurs. The reactorcontains a sufficient quantity of suitable catalyst to convertapproximately 50% of the dimethyl maleate to tetrahydrofuran. The streamfrom this reactor passes to the second vaporiser where it is used tovaporise refining recycle. The stream from the second vaporiser passesto the second reactor where conversion to crude product occurs. Theproducts from the second reactor are refined and the refining recyclesseparated and recycled to the second vaporiser. The selectivity totetrahydrofuran is measured. The reaction details and results are setout in Table 1. It can be seen that approximately 90% more dimethylmaleate is reacted than is possible with the procedure of the prior art.

EXAMPLE 3

[0088] In a process scheme in accordance with the present invention andas illustrated in FIG. 2, the compressor cycle gas stream is maintainedat the same absolute rate as that for Comparative Example 1. In thisarrangement, 1.5 kmol/h of dimethyl maleate is fed to the firstvaporiser. No feed dimethyl maleate is fed to the second vaporiser.However, 0.75 kmol/h of refining recycle is fed to the second vaporiser27. 311 kmol/h of hydrogen cycle gas and 7.6 kmol/h of make up hydrogenare fed to the first vaporiser to vaporise the dimethyl maleate feedbefore the vaporised stream passes to the first reactor where conversionto crude product occurs. The reactor contains a sufficient quantity ofsuitable catalyst to convert approximately 50% of the dimethyl maleateto tetrahydrofuran. The stream from this reactor passes to the secondvaporiser where it is used to vaporise refining recycle. The stream fromthe second vaporiser passes to the second reactor where conversion tocrude product occurs. The products from the second reactor are refinedand the refining recycles separated and recycled to the secondvaporiser. The selectivity to tetrahydrofuran is measured. The reactiondetails and results are set out in Table 1. It can be seen thatapproximately 110% more dimethyl maleate is reacted than is possiblewith the procedure of the prior art. TABLE 1 Comp. Example ExampleExample E.g. 1 1 2 3 Vap1 DMM feed kmol/h 1.0 1.5 1.5 1.5 Vap2 DMM Feedkmol/h N/A 0 0.4 0.6 Vap1 refining kmol/h 0.4 0 0 0 recycle Vap2refining kmol/h N/A 0.3 0.4 0.6 recycle Total DMM feed kmol/h 1.0 1.51.9 2.1 Increase in DMM % N/A 50 90 110 processed compared to Comp. e.g.1 Loop Pressure Bara 62 62 62 62 (exit final Reactor) Vap1 Temp ° C. 167168 169 167 Vap2 Temp ° C. N/A 192 187 182 Make up kmol/h 4.9 7.6 9.411.0 Hydrogen Cycle gas (at kmol/h 311 311 311 311 compressor) Cyclegas/DMM kmol/ 311 211 167 148 feed kmol Reactor 1 THF % 2.8 47.4 47.490.0 selectivity Overall THF % 2.8 47.0 47.0 61.8 selectivity

[0089] The dew points at various points in the first reaction zone aredetermined and compared. The results are set out in Table 2. Theseresults assume that the hydrogenation reaction is 100% and is followedby dehydration. In reality it will be understood that some, i.e. lessthan 10 mol % dehydration may occur in the hydrogenation zone and thatthere will be some hydrogenation of residual ester and/or lactone in thedehydration zone. It is also necessary to note that the system isnon-ideal and that it is necessary to allow for vapour pressure errorsand heat of reaction errors. TABLE 2 Comp E.g. 1 Example 1 Example 2Example 3 Reactor 1 ° C. 172 169 169 169 Inlet Temp Reactor 1 ° C. 161164 164 164 Inlet Dew Point Reactor 1 ° C. 190 195 195 195 Temp Exit -Hydrogenation Zone Reactor 1 ° C. 185 191 191 191 Dew Point ExitHydrogenation Zone Reactor 1 ° C. — 195 195 195 Temp Exit DehydrationZone Reactor 1 ° C. — 172 172 128 Dew Point Exit Dehydration ZoneReactor 1 ° C. 5 23 23 67 Exit Margin

[0090] Thus it will be understood that Examples 1 to 3 have asignificantly larger dew point margin exit for the first reactor thanfor the comparative example. This increase in dew point margin occursprimarily as a result of the dehydration of the butane-1,4-diol to thetetrahydrofuran that has taken place within the first reactor. It willbe noted that the wider the dew point margin exit from the firstreactor, the more feed material can be vaporised in the downstreamvaporiser with a corresponding increase in hydrogenation loopproductivity and reduction in the cycle gas flow per unit ofdimethylmaleate feed.

1. A process for the production of an ether by reaction of acorresponding organic feed material selected from dicarboxylic acidsand/or anhydrides, monoesters of dicarboxylic acids and/or anhydrides,diesters of dicarboxylic acids and/or anhydrides, lactones, and mixturesof two or more thereof in the presence of hydrogen which comprises thesteps of: (a) supplying a stream comprising the organic feed material toa first vaporization zone and contacting said feed with cycle gascomprising hydrogen such that at least a portion of the feed material isvaporized by and into the cycle gas; (b) supplying at least a portion ofthe cycle gas and the vaporized feed material to a first reaction zonecomprising catalyst and operating under reaction conditions to allowhydrogenation and dehydration to occur; (c) recovering from the firstreaction zone an intermediate product stream comprising unreacted feedmaterial, cycle gas, desired product(s), and any co-products andby-products; (d) supplying the intermediate product stream to a secondvaporization zone and contacting it with additional feed material suchthat the said additional feed material is vaporized by and into theintermediate product stream; (e) supplying the product of step (d) to asubsequent reaction zone comprising catalyst and operating underreaction conditions to allow hydrogenation and, if required, dehydrationto occur; and (f) recovering from the subsequent reaction zone a productstream comprising the ether.
 2. A process according to claim 1 whereinthe cycle gas and vaporized organic feed from step (a) is divided with amajor portion being supplied to step (b) and a minor portion to step(d).
 3. A process according to claim 1 wherein all of the cycle gas andthe vaporized feed material from step (a) is supplied to step (b).
 4. Aprocess according to claim 1 wherein the process additionally includesthe step of separating any co-products and/or by-products from theproduct stream in a refining zone and recycling one or more of saidco-products and/or by products in one or more recycle streams to one ormore of the vaporisation zones where they will be vaporised.
 5. Aprocess according to claim 4 wherein the organic feed stream fed to atleast one vaporization zone is, or includes, one or more recyclestreams.
 6. A process according to claim 1 wherein the process includesone or more additional subsequent reaction zones located in seriesbetween the first and final subsequent reaction zones and wherein the oreach additional subsequent reaction zone is preceded by a vaporizationzone in which additional feed, recycle or fresh feed and recycle arevaporized by and into the intermediate product stream from the previousreaction zone.
 7. A process according to claim 1 wherein the recyclestreams are vaporized into an intermediate product stream before beingmixed with cycle gas comprising additional vaporized feed material.
 8. Aprocess according to claim 1 wherein the organic feed material iscontained within an organic solvent which is separated from the feedmaterial by cycle gas stripping in one or more of the vaporizationzones.
 9. A process according to claim 1 wherein the catalyst is acombination of different catalysts selected from high acid tolerancecatalysts, high ester conversion hydrogenation catalysts and high etherformation catalysts.
 10. A process according to claim 1, wherein theoverall selectivity to the ether is more than 10%
 11. A processaccording to claim 1 wherein the selectivity to the ether is more than30% in at least one reaction zone.
 12. A process according to claim 1wherein the organic feed material is selected from mono- C₁ to C₄ alkyesters of C₄ to C₁₂ dicarboxylic acids and/or anhydrides, di- C₁ to C₄alkyl esters of C₄ to C₁₂ dicarboxylic acids and/or anhydrides, lactonesof C₄ to C₁₂ hydroxycarboxylic acids, and mixtures of two or morethereof.
 13. A process according to claim 12 wherein the organic feedmaterial is selected from monomethyl maleate, monomethyl fumarate,monomethyl succinate, dimethyl maleate, dimethyl fumarate, dimethylsuccinate, γ-butyrolactone, monoethyl maleate, monoethyl fumarate,monoethyl succinate, diethyl maleate, diethyl fumarate, diethylsuccinate, γ-butyrolactone, and mixtures of two or more thereof.
 14. Aprocess according to claim 1 wherein the ether is a cyclic ether.
 15. Aprocess according to claim 14 wherein the ether is tetrahydrofuran.